Apparatus for the reduction of gasoline benzene content by alkylation with dilute ethylene

ABSTRACT

The apparatus converts ethylene in a dilute ethylene stream and dilute benzene in an aromatic containing stream via alkylation to heavier hydrocarbons. The catalyst may be a zeolite such as UZM-8. The catalyst is resistant to feed impurities such as hydrogen sulfide, carbon oxides, and hydrogen and selectively converts benzene. At least 40 wt-% of the ethylene in the dilute ethylene stream and at least 20 wt-% of the benzene in the dilute benzene stream can be converted to heavier hydrocarbons.

CROSS-REFERENCE TO RELATED APPLICATION

This application is a Continuation of copending application Ser. No.12/813,534 filed Jun. 11, 2010, the contents of which are herebyincorporated by reference in its entirety.

BACKGROUND OF THE INVENTION

The field of the invention is an apparatus for alkylating benzene in anaphtha stream with dilute ethylene. The alkylated product may be usedas motor fuel.

Dry gas is the common name for the off-gas stream from a fluid catalyticcracking unit that contains all the gases with boiling points of ethaneand lower. The off-gas stream is compressed to remove as much of the C₃and C₄ gases as possible. Sulfur is also largely absorbed from theoff-gas stream in a scrubber that utilizes an amine absorbent. Theremaining stream is known as the FCC dry gas. A typical dry gas streamcontains 5 to 50 wt-% ethylene, 10 to 20 wt-% ethane, 5 to 20 wt-%hydrogen, 5 to 20 wt-% nitrogen, about 0.05 to about 5.0 wt-% of carbonmonoxide, 0.1 to about 5.0 wt-% of carbon dioxide and less than 0.01wt-% hydrogen sulfide and ammonia with the balance being methane.

Currently, the FCC dry gas stream is sent to a burner as fuel gas. AnFCC unit that processes 7,949 kiloliters (50,000 barrels) per day willburn about 181,000 kg (200 tons) of dry gas containing, for example,about 36,000 kg (40 tons) of ethylene as fuel per day. Because a largeprice difference exists between fuel gas and motor fuel products or pureethylene it would appear economically advantageous to attempt to recoverthis ethylene. However, the dry gas stream contains impurities that canpoison catalysts and is so dilute that ethylene recovery is noteconomically justified by gas recovery systems.

There is need for utilization of dilute ethylene in refinery streams.

Catalytic reforming is a well-established hydrocarbon conversion processemployed in the petroleum refining industry for improving the octanequality of hydrocarbon feedstocks, the primary product of reformingbeing motor gasoline. In catalytic reforming, a naphtha feedstock isadmixed with a recycle stream comprising hydrogen and contacted withcatalyst in a reaction zone at temperatures of around 493° to 510° C.(920° to 950° F.) and moderate pressure around 1379 to 3792 kPa (200 to550 psig). The usual feedstock for catalytic reforming is a petroleumfraction known as naphtha and having an initial boiling point of about46° C. (115° F.) and an end boiling point of about 204° C. (400° F.).

The catalytic reforming process is particularly applicable to thetreatment of straight run gasoline comprised of relatively largeconcentrations of naphthenic and substantially straight chain paraffinichydrocarbons, which are subject to aromatization through dehydrogenationand/or cyclization reactions. The catalyst “reforms” the molecularstructures of the hydrocarbons contained in the raw naphtha by removinghydrogen and rearranging the structure of the molecules so as to improvethe octane number of the naphtha. However, the increase in octane numberalso reduces the liquid volume of the naphtha as the specific gravity isincreased. Because of the multiplicity of the compounds in the rawnaphtha, the actual reactions which occur in catalytic reforming arenumerous. However, some of the many resulting products are aryl oraromatic compounds, all of which exhibit high octane numbers. The arylcompounds produced depend upon the starting materials which in arefinery are controlled by the boiling range of the naphtha used and thecrude oil source. The “reformed” product from a catalytic reformingprocess is commonly called reformate and is often separated into twofractions by conventional distillations, a light reformate having aboiling range of about 46° to 121° C. (115° to 250° F.) and a heavyreformate having a boiling range of about 121° to 204° C. (250° to 400°F.). The aryl compounds in each fraction are thus dependent upon theirboiling points. The lower boiling or lighter aryl compounds, e.g.,benzene, toluene and xylenes, are contained in the light reformate, andhigher boiling aryl compounds are contained in the heavy reformate.

The concentration of benzene in gasoline is now being regulated by theAmerican government. The Mobil Source Air Toxics regulation (MSAT II)requires that the average benzene level in gasoline produced by arefiner be lower than 0.62 vol-% with a maximum of 1.3 vol-% in gasolineproduced at any one refinery. Benzene is commonly produced at levelshigher than this by reforming processes and FCC processes. As reformateand the naphtha streams from the FCC unit are two of the largest sourcesof gasoline in a refinery, benzene reduction strategies have to be used.

Currently, benzene is commonly sent to a saturation unit to reducebenzene to cyclohexane. However, this process utilizes at least threemoles of hydrogen for every mole of benzene converted and there is anoctane loss associated with the conversion of benzene to cyclohexane.Methods for the reduction of benzene in gasoline without loss of octaneor use of hydrogen are necessary.

The alkylation of concentrated benzene streams with concentratedethylene streams is known. Alkylation typically involves the use ofclean ethylene streams because alkylation catalysts are susceptible tofeed impurities. Additionally, dilute ethylene is little used as anoligomerization feedstock because of its much lower reactivity relativeto higher olefins. Benzene streams fed to alkylation reactors are alsoconcentrated because of concern that heavier aromatics willpreferentially alkylate, thereby requiring the use of a large excess ofethylene before reducing the benzene concentration and producingundesirable polyalkylated benzene.

DEFINITIONS

The following definitions shall be applicable throughout this document.

The term “communication” means that material flow is operativelypermitted between enumerated components.

The term “downstream communication” means that at least a portion ofmaterial flowing to the subject in downstream communication mayoperatively flow from the object with which it communicates.

The term “upstream communication” means that at least a portion of thematerial flowing from the subject in upstream communication mayoperatively flow to the object with which it communicates.

The term “column” means a distillation column or columns for separatingone or more components of different volatilities based on boiling pointdifferential. A column may have a reboiler on its bottom and a condenseron its overhead. Unless otherwise indicated, each column includes acondenser on an overhead of the column to condense and reflux a portionof an overhead stream back to the top of the column and a reboiler at abottom of the column to vaporize and send a portion of a bottoms streamback to the bottom of the column. Feed to a column may be preheated. Thetop pressure is the pressure of the overhead vapor at the outlet of thecolumn. The bottom temperature is the liquid bottom outlet temperature.Overhead lines and bottoms lines refer to the net lines from the columndownstream of the reflux or reboil to the column.

As used herein, the term “a component-rich stream” means that the richstream coming out of a vessel has a greater concentration of thecomponent than the feed to the vessel.

As used herein, the term “a component-lean stream” means that the leanstream coming out of a vessel has a smaller concentration of thecomponent than the feed to the vessel.

The term “C_(x)” is to be understood to refer to molecules having thenumber of carbon atoms represented by the subscript “x”. Similarly, theterm “C_(x)−” refers to molecules that contain less than or equal to xand preferably x and less carbon atoms. The term “C_(x)+” refers tohydrocarbons with more than or equal to x and preferably x and morecarbon atoms.

SUMMARY OF THE INVENTION

We have found that dilute benzene in aromatic containing streams such asreformate or FCC light naphtha can be alkylated over zeolitic catalystswith ethylene in dilute ethylene streams, such as an FCC dry gas stream.The heavier hydrocarbons can be separated and blended in the gasolineand diesel pools. We have found most zeolitic catalysts that aresuitable for alkylation of benzene with light olefins quickly deactivatein dilute ethylene streams. Neither the dilute nature of the ethylene,nor the impurities present substantially affect a catalyst comprisingUZM-8. Additionally, conversion of benzene by alkylation with ethyleneis as high as toluene and greater than heavier aromatics with UZM-8catalyst. Consequently, dilute benzene in a naphtha stream such asreformate can be alkylated with dilute ethylene in an FCC dry gas streamto provide a liquid fuel product which is reduced in benzeneconcentration and easy to separate from the unconverted gas stream. Theunconverted gas can then be burned as fuel gas, but with the morevaluable ethylene recovered as gasoline range hydrocarbons.

Advantageously, the apparatus can enable utilization of dilute ethylenein a stream and in the presence of feed impurities that can be catalystpoisons.

Advantageously, the apparatus can enable reduction of the concentrationof dilute benzene in a naphtha stream without utilizing hydrogen orreducing the liquid volume or the octane value of the naphtha streameven in the presence of heavier aromatic hydrocarbons.

In an embodiment, the invention comprises an apparatus for alkylatingbenzene with ethylene comprising a fluid catalytic cracking reactor forcontacting cracking catalyst with a hydrocarbon feed stream to crack thehydrocarbon feed to cracked products having lower molecular weight anddeposit coke on the cracking catalyst to provide coked crackingcatalyst. The apparatus also comprises a regenerator for combusting cokefrom the coked cracking catalyst by contact with oxygen and a separatorin communication with the fluid catalytic cracking reactor forseparating C₃ hydrocarbons from C₂ hydrocarbons to provide a diluteethylene stream. A reforming reactor for contacting a naphtha streamwith reforming catalyst to produce a reformate stream is also comprisedin the apparatus. Lastly, the apparatus comprises an alkylation reactorin communication with the separator and the reforming reactor foralkylating benzene in the reformate stream with ethylene in the diluteethylene stream over a fixed bed of alkylation catalyst to heavier alkylbenzene hydrocarbons. No fractionation column is in communicationbetween the reforming reactor and the alkylation reactor. In analternative embodiment, no fractionation column is in communicationbetween the separator and the alkylation reactor. In a still furtherembodiment, an absorber in communication with the product outlet of theFCC reactor provides an off-gas stream comprising a dilute ethylenestream and no fractionation column is in communication between thereforming reactor and the alkylation reactor and between the absorberand the alkylation reactor.

Additional features and advantages of the invention will be apparentfrom the description of the invention, the drawings and claims providedherein.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic drawing of an FCC unit, a reforming unit and analkylation unit.

FIG. 2 is a graph of conversion and temperature over time.

FIG. 3 is a graph of conversion over time.

DETAILED DESCRIPTION

The present invention may be applied to any hydrocarbon streamcontaining ethylene and, preferably, a dilute proportion of ethylene. Asuitable, dilute ethylene stream may typically comprise between about 5and about 50 wt-% ethylene. An FCC dry gas stream is a suitable diluteethylene stream. Other dilute ethylene streams may also be utilized inthe present invention such as coker dry gas streams. Because the presentinvention is particularly suited to FCC dry gas, the subject applicationwill be described with respect to utilizing ethylene from an FCC dry gasstream.

The present invention may also be applied to any hydrocarbon streamcontaining benzene and, preferably, a dilute proportion of benzene. Asuitable benzene stream may typically comprise between about 1 and about50 wt-% benzene, at least about 3 wt-% toluene and at least about 20wt-% paraffins. A reformate stream is a suitable benzene stream. Otherbenzene streams may also be utilized in the present invention such asFCC aromatic naphtha streams. Because the present invention isparticularly suited to a reformate stream, the subject application willbe described with respect to utilizing benzene from a reformate stream.

Now turning to FIG. 1, wherein like numerals designate like components,the FIG. 1 illustrates a refinery complex 6 that generally includes anFCC unit 10 with a product recovery section 90, a reformer unit 200 andan alkylation unit 300.

The FCC unit 10 includes a reactor 12 and a catalyst regenerator 14.Process variables typically include a cracking reaction temperature of400° to 600° C. and a catalyst regeneration temperature of 500° to 900°C. Both the cracking and regeneration occur at an absolute pressurebelow 506 kPa (72.5 psia).

FIG. 1 shows a typical FCC reactor 12 in which a heavy hydrocarbon feedor raw oil stream 16 is contacted from a distributor with a regeneratedcracking catalyst entering from a regenerated catalyst standpipe 18.This contacting may occur in a narrow riser 20, extending upwardly tothe bottom of a reactor vessel 22. The contacting of feed and catalystis fluidized by gas from a fluidizing line 24. In an embodiment, heatfrom the catalyst vaporizes the hydrocarbon feed or oil, and thehydrocarbon feed is thereafter cracked to lighter molecular weighthydrocarbon products in the presence of the catalyst as both aretransferred up the riser 20 into the reactor vessel 22. Inevitable sidereactions occur in the riser 20 leaving coke deposits on the catalystthat lower catalyst activity. The cracked light hydrocarbon products arethereafter separated from the coked cracking catalyst using cyclonicseparators which may include a primary separator 26 and one or twostages of cyclones 28 in the reactor vessel 22. Gaseous, crackedproducts exit the reactor vessel 22 through a product outlet 31 to line32 for transport to a downstream product recovery section 90. The spentor coked catalyst requires regeneration for further use. Coked crackingcatalyst, after separation from the gaseous product hydrocarbons, fallsinto a stripping section 34 in which steam is injected through a nozzleto purge any residual hydrocarbon vapor. After the stripping operation,the coked catalyst is carried to the catalyst regenerator 14 through aspent catalyst standpipe 36.

FIG. 1 depicts a regenerator 14 known as a combustor. However, othertypes of regenerators are suitable. In the catalyst regenerator 14, astream of oxygen-containing gas 30, such as air, is introduced throughan air distributor 38 to contact the coked catalyst. Coke is combustedfrom the coked catalyst to provide regenerated catalyst and flue gas.The catalyst regeneration process adds a substantial amount of heat tothe catalyst, providing energy to offset the endothermic crackingreactions occurring in the reactor riser 20. Catalyst and air flowupwardly together along a combustor riser 40 located within the catalystregenerator 14 and, after regeneration, are initially separated bydischarge through a disengager 42. Additional recovery of theregenerated catalyst and flue gas exiting the disengager 42 is achievedusing first and second stage separator cyclones 44, 46, respectivelywithin the catalyst regenerator 14. Catalyst separated from flue gasdispenses through diplegs from cyclones 44, 46 while hot flue gasrelatively lighter in catalyst sequentially exits cyclones 44, 46 andexits the regenerator vessel 14 through flue gas outlet 47 in flue gasline 48. Regenerated catalyst is carried back to the riser 20 throughthe regenerated catalyst standpipe 18. As a result of the coke burning,the flue gas vapors exiting at the top of the catalyst regenerator 14 inline 48 contain CO, CO₂, N₂ and H₂O, along with smaller amounts of otherspecies.

The product recovery section 90 is in downstream communication with theproduct outlet 31. In the product recovery section 90, the gaseous FCCproduct in line 32 is directed to a lower section of an FCC mainfractionation column 92. The main column 92 is in downstreamcommunication with the product outlet 31. Several fractions of FCCproduct may be separated and taken from the main column including aheavy slurry oil stream from the bottoms in line 93, a heavy cycle oilstream in line 94, a light cycle oil in line 95 taken from outlet 95 aand a heavy naphtha stream in line 96 taken from outlet 96 a. Any or allof lines 93-96 may be cooled and pumped back to the main column 92 tocool the main column typically at a higher location. Gasoline andgaseous light hydrocarbons are removed in overhead line 97 from the maincolumn 92 and condensed before entering a main column receiver 99. Themain column receiver 99 is in downstream communication with the productoutlet 31, and the main column 92 is in upstream communication with themain column receiver 99.

An aqueous stream is removed from a boot in the receiver 99. Moreover, acondensed light naphtha stream is removed in line 101 while an overheadstream is removed in line 102. The overhead stream in line 102 containsgaseous light hydrocarbon which may comprise a dilute ethylene stream. Aportion of the condensed light naphtha stream in line 101 may berefluxed back to the main fractionation column 92 in line 105 leaving anet condensed light naphtha stream in line 103. The streams in lines 101and 102 may enter a vapor recovery section 120 of the product recoverysection 90.

The vapor recovery section 120 is shown to be an absorption basedsystem, but any vapor recovery system may be used including a cold boxsystem. To obtain sufficient separation of light gas components, thegaseous stream in line 102 is compressed in compressor 104. More thanone compressor stage may be used, and typically a dual stage compressionis utilized to compress the gaseous stream in line 102 to between about1.2 MPa to about 2.1 MPa (gauge) (180 to 300 psig). Three stages ofcompression may be advantageous to provide additional pressure at leastas high as 3.4 MPa (gauge) (500 psig).

The compressed light hydrocarbon stream in line 106 is joined by C₃+hydrocarbon liquid stream in a primary absorber bottoms line 107 and aC₂− hydrocarbon stream in stripper overhead line 108, chilled anddelivered to a high pressure receiver 110. An aqueous stream from thereceiver 110 may be routed to the main column receiver 99. A gaseousfirst FCC product stream in line 112 comprising a dilute ethylene streamis routed to a unit that effects a separation between C₃+ and C₂−hydrocarbons, which in this embodiment is a primary absorber 114. In theprimary absorber 114 the dilute ethylene, first FCC product stream iscontacted with a second FCC product stream comprising unstabilizedgasoline from the main column receiver 99 in line 103 to effect aseparation between C₃+ and C₂− hydrocarbons. The separator forseparating C₃ from C₂ hydrocarbons, which may be the primary absorber114, is in downstream communication with the main column receiver 99. Aliquid C₃+ stream in line 107 is returned to line 106 prior to chilling.An overhead of the primary absorber 114 comprising dry gas ofpredominantly C₂− hydrocarbons with hydrogen sulfide, ammonia, carbonoxides and hydrogen is removed in the primary off-gas stream in line 116to comprise a dilute ethylene stream. However, to concentrate theethylene stream further and to recover heavier components line 116 mayoptionally be directed to a second unit that effects a separationbetween C₃+ and C₂− hydrocarbons, which in this embodiment is asecondary absorber 118. In the secondary absorber, a circulating streamof light cycle oil in line 121 diverted from line 95 absorbs most of theremaining C₅+ and some C₃-C₄ material in the primary off-gas stream. Thesecondary absorber 118 is in downstream communication with the primaryabsorber 114. Light cycle oil from the bottom of the secondary absorberin line 119 richer in C₃+ material is returned to the main column 92 viathe pump-around for line 95. An overhead of the secondary absorber 118comprising dry gas of predominantly C₂− hydrocarbons with hydrogensulfide, ammonia, carbon oxides and hydrogen is removed in the secondaryoff-gas stream in line 122 to comprise a dilute ethylene stream. Both ofthe absorber columns 114 and 118 have no condenser or reboiler, but mayemploy pump-around cooling circuits.

Liquid from the high pressure receiver 110 in line 124 is sent to astripper column 126. Most of the C₂− is removed in the overhead of thestripper 126 and returned to line 106 via overhead line 108. Thestripper column 126 has no condenser but receives cooled liquid feed inline 124. A liquid bottoms stream from the stripper 126 is sent to adebutanizer column 130 via line 128. An overhead stream in line 132 fromthe debutanizer comprises C₃-C₄ olefinic product while a bottoms streamin line 134 comprising stabilized gasoline may be further treated andsent to gasoline storage. In an embodiment, the bottoms stream in line134 may be sent to a naphtha splitter column 140. Light naphtha may berecovered in an overhead line 142 from the naphtha splitter column 140with a heavy naphtha stream recovered in bottoms line 144. An aromaticnaphtha comprising benzene may be taken as a mid cut in line 146 inwhich case the naphtha splitter column 140 may be a dividing wallcolumn. One or both of light and heavy naphtha streams in lines 142 and144 may be taken to a gasoline distribution system 148. Only heavynaphtha stream in line 144 is shown going to gasoline distributionsystem 148 in FIG. 1 because the light naphtha stream in line 142 may bedirected to processing separate from the heavy naphtha.

The dilute ethylene stream of the present invention may comprise an FCCdry gas stream comprising between about 5 and about 50 wt-% ethylene andpreferably about 10 to about 30 wt-% ethylene. Methane will typically bethe predominant component in the dilute ethylene stream at aconcentration of between about 25 and about 55 wt-% with ethane beingsubstantially present at typically between about 5 and about 45 wt-%.Between about 1 and about 25 wt-% and typically about 5 to about 20 wt-%of hydrogen and nitrogen each may be present in the dilute ethylenestream. Saturation levels of water may also be present in the diluteethylene stream. The dilute ethylene stream in overhead line 116 mayhave no more than 3 wt-% and suitably no more than 1 wt-% propylene andtypically no more than 25 wt-% and suitably no more than 15 wt-% C₃+materials. If the secondary absorber 118 is used, the dilute ethylenestream in overhead line 122 may have no more than about 5 wt-% of C₃+with typically less than 0.5 wt-% propylene. Besides hydrogen, otherimpurities such as hydrogen sulfide, ammonia, carbon oxides andacetylene may also be present in the dilute ethylene stream.

Impurities in a dry gas ethylene stream can poison an alkylationcatalyst. Carbon dioxide and ammonia can attack acid sites on thecatalyst. Hydrogen sulfide is known at times to deactivate zeolitecatalyst. Acetylene can polymerize and gum up on the catalyst orequipment.

The primary off-gas stream in line 116 or secondary off-gas stream inline 122, comprising a dilute ethylene stream may be introduced into anoptional amine absorber unit 150 to remove hydrogen sulfide to lowerconcentrations. An aqueous amine solution, such as comprisingmonoethanol amine or diethanol amine, is introduced via line 152 intoabsorber 150 and is contacted with the flowing off-gas stream to absorbhydrogen sulfide, and a hydrogen sulfide-rich aqueous amine absorptionsolution is removed from absorption zone 150 via line 154 and recoveredand perhaps further processed.

An optionally hydrogen sulfide-lean, amine-treated dilute ethylenestream in line 156 may be introduced into an optional water wash unit160 to remove residual amine carried over from the amine absorber 150and reduce the concentration of ammonia and carbon dioxide in the diluteethylene stream in line 156. Water is introduced to the water wash inline 162. The water in line 162 is typically slightly acidified toenhance capture of basic molecules such as the amine. An aqueous streamin line 164 rich in amine and potentially ammonia and carbon dioxideleaves the water wash unit 160 and may be further processed.

The optionally amine treated, optionally water washed dilute ethylenestream in line 166 may then be treated in an optional guard bed 170 toremove one or more of the impurities such as carbon monoxide, hydrogensulfide and ammonia down to lower concentrations. The guard bed 170 maycontain an adsorbent to adsorb impurities such as ammonia that maypoison an alkylation catalyst. The guard bed 170 may contain multipleadsorbents for adsorbing more than one type of impurity. A typicaladsorbent for adsorbing hydrogen sulfide is ADS-12, for adsorbing carbonmonoxide is ADS-106 and for adsorbing ammonia is UOP MOLSIV 3A allavailable from UOP, LLC. The adsorbents may be mixed in a single bed orcan be arranged in successive beds.

The optionally amine-treated, optionally water-washed and optionallyadsorption-treated stream in line 172 may be dried in a dryer 174 toremove water down to below about 500 wppm water. Water can adverselyaffect alkylation catalyst.

A dilute ethylene stream in line 176 optionally amine treated,optionally water washed, optionally adsorption treated and optionallydried will typically have at least one of the following impurityconcentrations: about 0.05 wt-% and up to about 5.0 wt-% of carbonmonoxide and/or about 0.1 wt-% and up to about 5.0 wt-% of carbondioxide, and/or at least about 1 wppm and up to about 500 wppm hydrogensulfide and/or at least about 1 and up to about 500 wppm ammonia, and/orat least about 5 and up to about 20 wt-% hydrogen. The type ofimpurities present and their concentrations will vary depending on theprocessing and origin of the dilute ethylene stream.

Line 176 carries the dilute ethylene stream to a compressor 180 ifnecessary to be pressured up to alkylation reactor pressure. Thecompressor 180 is in downstream communication with the main column 92,the product recovery section 90 and the product outlet 31. Thecompressor 180 may comprise one or more stages with interstage cooling.A heater may be required to bring the compressed stream up to reactiontemperature. The compressed dilute ethylene gas stream is carried inline 182 to the alkylation unit 300. Line 182 feeds the dilute ethylenestream to the alkylation reactor unit 300. The alkylation reactor unit300 may be in downstream communication with the compressor 180 and/orthe first or second separators for separating C₃ hydrocarbons from C₂hydrocarbons, which may be the primary or secondary absorbers 114 and118, respectively. In an embodiment, no fractionation unit is incommunication between the first separator or second separator forseparating C₃ hydrocarbons from C₂ hydrocarbons and the alkylationreactor 320. Consequently, in an embodiment no fractionation unit is incommunication between the primary absorber 114 or the secondary absorber118 and the alkylation reactor 320. In this embodiment, the diluteethylene stream may be subjected to separations based on adsorption orabsorption, but no fractionation based on boiling point differential isconducted on the dilute ethylene stream which may comprise primary orsecondary off-gas between the primary absorber 114 and/or secondaryabsorber 118 and the alkylation reactor 320. This embodiment stands incontradistinction to conventional belief which held that a diluteethylene dry gas stream required fractionation such as in a demethanizercolumn to remove lighter components before the ethylene could sustain analkylation reaction with benzene. The obviation of a demethanizer columnresults in substantial operating and capital savings.

Turning to the reforming unit 200, naphtha feedstock stream in line 202is admixed with a stream comprising hydrogen from line 204, heated andcontacted with catalyst in a reforming reactor 210 to produce areformate. Desirably, the reforming reactor 210 is a moving bed reactorthat receives regenerated catalyst through a line 220 and dischargesspent catalyst through a line 222 to a regeneration zone 230 assisted byfluidized inert gas. Catalyst flows from the top to the bottom of thestacked reactor 210, passing first through a reduction zone 224, inwhich a hydrogen-rich gas from line 226 contacts and reduces theoxidized catalyst particles. From there, catalyst flows through multiplereaction zones in which naphtha feed contacts the catalyst particles.The reforming reactor 210 can comprise a stacked reactor arrangement,which can include a plurality of reactions zones. Each reaction zone hasa catalyst bed in the stacked reactor 210 to permit continuous orintermittent flow of the catalyst particles from the top reaction zone212 through second and third reaction zones 214 and 216, respectively,to a final zone 218. Effluent from the first through third reactionzones 212-216 may be withdrawn, heated and returned to the subsequentreaction zone 214-218, respectively. More or less catalyst beds can beused. A reformate product stream may be withdrawn in line 232 from thefinal reaction zone 218. A lower retention chamber 234 at the bottom ofthe stacked reactor 210 receives spent catalyst. A purging fluidpreferably comprising hydrogen enters lower retention chamber 234 from aline 236 at a rate that purges hydrocarbons from the catalyst particlesin lower retention chamber 234.

The usual feedstock for catalytic reforming is a petroleum fractionknown as naphtha and having an initial boiling point of about 46° C.(115° F.) and an end boiling point of about 204° C. (400° F.). Thecatalytic reforming process is particularly applicable to the treatmentof straight run naphtha comprised of relatively large concentrations ofnaphthenic and substantially straight chain paraffinic hydrocarbons,which are subject to aromatization through dehydrogenation and/orcyclization reactions. In reforming, dehydrogenation of cyclohexanes anddehydroisomerization of alkylcyclopentanes yield aromatics,dehydrogenation of paraffins yields olefins, dehydrocyclization ofparaffins and olefins yields aromatics, isomerization of n-paraffins andalkylcycloparaffins yield cyclohexanes, substituted aromatics areisomerized and paraffins are hydrocracked.

A catalytic reforming reaction is normally effected in the presence ofcatalyst particles comprised of one or more Group VIII noble metals suchas platinum, iridium, rhodium, palladium, and a halogen combined with aporous carrier, such as a refractory inorganic oxide. The halogen isnormally chloride. Alumina is a commonly used carrier. The preferredalumina materials are known as the gamma, eta and theta alumina withgamma and eta alumina giving the best results. An important propertyrelated to the performance of the catalyst is the surface area of thecarrier. Preferably, the carrier will have a surface area of from 100 toabout 500 m²/g. The particles are usually spheroidal and have a diameterof from about 1/16th to about ⅛th inch (1.5-3.1 mm), though they may beas large as ¼th inch (6.35 mm). During the course of a reformingreaction, catalyst particles become deactivated as a result ofmechanisms such as the deposition of coke on the particles; that is,after a period of time in use, the ability of catalyst particles topromote reforming reactions decreases to the point that the catalyst isno longer useful. The catalyst must be reconditioned, or regenerated,before it can be reused in a reforming process.

The regeneration zone 230 regenerates catalyst and recycles it to line220. Spent catalyst particles containing coke deposits flow from thelower retention chamber 234 of the stacked reactor 210 through a liftconduit 222 into a disengaging vessel 240 in the regeneration zone 230.The disengaging vessel 240 may comprise two sections. In an uppersection 242, an elutriation fluid enters upper section 242 at a ratethat separates broken or chipped catalyst particles and catalyst finesfrom the whole catalyst particles which exit the bottom of thedisengaging vessel 240. The catalyst chips and fines pass out of theupper section 242 with the elutriation fluid which may be filtered andrecycled to the upper section 242. In the lower section 244, elutriatedcatalyst is contacted with cooled flue gas from line 246 to adsorbchlorines and hydrogen chloride onto the catalyst. The flue gas in line246 may be heat exchanged with dechlorinated flue gas in line 248 fromthe lower section 244 to cool it into an adsorptive condition before itis fed to the lower section 244. Catalyst passes from the lower section244 into the catalyst regenerator 250 via lines 252

The catalyst regenerator 250 comprises a combustion section 254 and aconditioning section 256. In the combustion section, catalyst descendsin an inner annular chamber 258 which may comprise concentric screensimpermeable to catalyst passage. Dechlorinated flue gas recycled in line248 is heated by heat exchange with flue gas in line 246 and fed to anouter annular chamber 260. Recycled flue gas passes from outer annularchamber 260 through an outer screen into the inner annular chamber 258to heat the catalyst therein. Gases used in the conditioning section 256comprising oxygen and chlorine ascending from the conditioning sectionalso passes into the inner annular chamber 258 to combust coke depositsfrom the catalyst. Hot combustion gases exit the inner concentric screeninto an inner pipe 245 and leave the combustion section 254 in line 246.Combusted catalyst descends into the conditioning chamber 256 into acentral space 262 within an annular baffle. Air which may be enrichedwith oxygen and mixed with chlorine is fed in line 264 to a lowerannular chamber 266 outside of the annular baffle and passes into thecentral space 262 in contact with the catalyst to disperse the metal onthe catalyst. Regenerated catalyst exits the conditioning section 256and is lifted in line 220 back to the reforming reactor 210 assisted byfluidized gas.

A reformate splitter column 270 may be in downstream communication withthe reforming reactor 210 and in communication between the reformingreactor 210 and an alkylation reactor via line 232. In this embodiment,the control valve on line 278 is at least partially closed, and thecontrol valve on line 272 is at least partially open so line 272 is ableto feed the reformate product stream from line 232 to reformate splittercolumn 270. Fractionation produces a light reformate stream in overheadline 274 which then feeds a benzene stream through an open control valvethereon into line 304 and alkylation reactor 320. The light reformatebenzene stream can have about 20 to about 50 wt-% benzene, with at least20 wt-% paraffins and the balance of at least 3 wt-% toluenes. In anembodiment, the light reformate benzene stream has a greaterconcentration of paraffins than benzene. The benzene stream preferablycomprises at least about 4.0 wt-% benzene. The bottoms stream of heavyreformate in line 276 exiting a bottom half of the reformate splittercolumn 270 may bypass the alkylation reactor 320 and be routed to thegasoline distribution system 148 without communicating with thealkylation reactor 320.

Alternatively, the full reformate stream in line 232 may bypass thereformate splitter 270 or the reforming unit 200 may omit a reformatesplitter column 270 altogether. In this embodiment, the control valve onlines 272 and 274 are closed, and the control valve on line 278 is atleast partially open, so line 232 is able to feed the full reformateproduct stream to line 304 and the alkylation reactor 320. In thisembodiment, the full reformate stream may be directed to an alkylationunit 300 in line 278. The reformate splitter may be omitted entirely. Inthis embodiment, no fractionation column is in communication between thereforming reactor 210 and the alkylation reactor 320. In thisembodiment, the full reformate stream may be subjected to separationsbased on adsorption or absorption, but no fractionation based on boilingpoint differential is conducted on the reformate stream between thereforming reactor 210 and the alkylation reactor 320. This embodimentstands in contradistinction to conventional belief which held that afull reformate stream required fractionation such as in a reformatesplitter column to remove heavy aromatics before the benzene couldsustain an alkylation reaction with olefins. The obviation of areformate splitter column results in substantial operating and capitalsavings.

The full reformate stream will comprise about 1 to about 10 wt-%benzene, about 3 to about 30 wt-% toluenes with the balance being atleast about 20 wt-% paraffins and heavier aromatics. The full reformatestream comprises a greater concentration of aromatics with a molecularweight larger than benzene than a concentration of benzene.Specifically, the full reformate stream may have a greater concentrationof aromatics with eight carbon atoms than a concentration of benzene.Additionally, the full reformate stream may have a greater concentrationof paraffins than benzene.

Line 304 feeds the benzene stream to the alkylation unit 300. Thealkylation unit 300 may be in downstream communication with thereformate splitter column 270. An alkylation reactor 320 preferablycontains a fixed catalyst bed 342 and may contain a plurality ofcatalyst beds 342-46. The catalyst is preferably an UZM-8 zeolite boundwith alumina.

In an embodiment, diolefins in the bottoms stream in the gaseous diluteethylene stream in line 182 may optionally be delivered to be firstreacted with a selective hydrogenation catalyst in selectivehydrogenation zone 310, to selectively saturate diolefins withoutcompletely saturating them to paraffins. Suitable conditions foroperation of a selective hydrogenation process include passing thedilute ethylene stream in line 302 in the gas phase and hydrogen fromline 308 at molar ratio 0.5 to 5 moles hydrogen per mole of diolefinover a catalyst comprising at least one metal selected from the groupformed by nickel, palladium and platinum, deposited on a support such asaluminum oxide, at a temperature of 20° to 200° C. (68° to 392° F.), apressure of 689 to 3447 kPa(g) (100 to 500 psig), and a space velocityof 0.5 to 10 hr⁻¹. Two or more reaction zones may be used although onlyone is shown. Each reaction zone may employ a recycle (not shown) ofreactor effluent to the reactor inlet with a ratio of recycle toethylene feed stream ranging from 0 to 20. The residual diolefin contentof such a process can be in the range 1 to 100 wppm, depending on theseverity of the operation.

A dilute ethylene stream from selective hydrogenation reactor 310 inline 312 is injected into an alkylation reactor 320. In an aspect, adrier (not shown) on line 312 may be used to remove water to lowconcentrations which could affect the alkylation catalyst. Other guardbeds are also contemplated to remove catalyst poisons such as removingammonia or amines down to about 1 to about 500 wppm. One of the catalystbeds may serve as a guard bed to remove water and catalyst poisons. If adrier or a guard bed is used on the ethylene feed side in the alkylationunit 300, then one, part or both of drier 176 and adsorbent bed 170 maybe omitted from the product recovery section 90. The dilute ethylenestream in line 312 has essentially the same composition as in line 114or 122 with the exception of removed impurities. The dilute ethylenestream in line 312 optionally amine treated, optionally water washed,optionally adsorption treated and optionally dried may typically have atleast one of the following impurity concentrations: about 0.05 wt-% andup to about 5.0 wt-% of carbon monoxide and/or about 0.1 wt-% and up toabout 5.0 wt-% of carbon dioxide, and/or at least about 1 wppm and up toabout 500 wppm hydrogen sulfide and/or at least about 1 and up to about500 wppm ammonia and amines, and/or at least about 5 and up to about 20wt-% hydrogen.

Although transalkylation reactions may occur in the alkylation reactor320, alkylation reactions are predominant. The alkylation reactor isshown as an upflow reactor, but a downflow reactor may also be suitable.The stream in line 312 of dilute ethylene is injected into thealkylation reactor 320 in several lines 322, 324 and 326 into pre-bedspaces 332, 334 and 336 prior to entry into catalyst beds 342, 344 and346, respectively. The catalyst beds 342, 344 and 346 contain alkylationcatalyst to alkylate ethylene onto benzene to produce ethylbenzene andonto toluene to produce ethyltoluene. Other alkylation reactions occurto produce alkylbenzenes and alkylaromatics. The liquid benzene streamin line 304 is fed to the alkylation reactor 320 where it initiallyabsorbs the dilute ethylene stream from the line 326 in the pre-bedspace 336 and together enter the catalyst bed 346. The aromaticreformate feed stream in line 304 may also receive the FCC aromaticstream from line 146 before entering the alkylation reactor 320.Alternatively, if the naphtha splitter only provides two streams, one ofthese may feed line 304 with aromatic naphtha which is preferably lightnaphtha stream in line 142. Gaseous ethylene dissolves into the liquidreformate stream to alkylate with aromatic rings.

The effluent from the catalyst bed 346 is mixed with fresh dry gascomprising dilute ethylene from the line 324 in the pre-bed space 334and together enter into the catalyst bed 344. The effluent from thecatalyst bed 344 is mixed with fresh dry gas comprising dilute ethylenefrom the line 322 in the pre-bed space 332 and together enter into thecatalyst bed 342. The process is repeated for the number of beds in thealkylation reactor 300. Although three catalyst beds are shown in thealkylation reactor 300, more or less beds and additional reactors may besuitable. Alkylation effluent from the alkylation reactor 320 istransported in an effluent line 352. A heat exchanger 354 may cool theeffluent in the line 352 to a desirable temperature. The alkylationreactor effluent stream may be depressurized by passing through apressure control valve or pressurized by passing through a pump, neitherof which is shown.

A preferred alkylation catalyst of the present invention is described asfollows. The preferred alkylation catalyst comprises an UZM-8 zeolite.One of the components of the catalyst support utilized in the presentinvention is alumina. The alumina source may be any of the varioushydrous aluminum oxides or alumina gels such as alpha-aluminamonohydrate of the boehmite or pseudo-boehmite structure, alpha-aluminatrihydrate of the gibbsite structure, beta-alumina trihydrate of thebayerite structure, and the like. A particularly preferred alumina isavailable from Sasol North America Alumina Product Group under thetrademark Catapal. This material is an extremely high purityalpha-alumina monohydrate (pseudo-boehmite) which after calcination at ahigh temperature has been shown to yield a high purity gamma-alumina.The zeolitic component of the catalyst is UZM-8 described in U.S. Pat.No. 6,756,030.

A suitable alkylation catalyst is prepared by mixing proportionatevolumes of UZM-8 and alumina to achieve the desired zeolite-to-aluminaratio. In an embodiment, 70 wt-% UZM-8 and 30 wt-% alumina powder willprovide a suitable support. In an embodiment, weight ratios other than70-to-30 of UZM-8 to alumina may be suitable, ranging from 90 wt-% UZM-8in content to 20 wt-% UZM-8 content with the balance alumina.

Monoprotic acid such as nitric acid or formic acid may be added to themixture in aqueous solution to peptize the alumina in the binder.Additional water may be added to the mixture to provide sufficientwetness to constitute a dough with sufficient consistency to be extrudedor spray dried.

The paste or dough may be prepared in the form of shaped particulates,with the preferred method being to extrude the dough through a diehaving openings therein of desired size and shape, after which theextruded matter is broken into extrudates of desired length and dried. Afurther step of calcination may be employed to give added strength tothe extrudate. Generally, calcination is conducted in a stream of dryair at a temperature from about 260° C. (500° F.) to about 815° C.(1500° F.).

The extruded particles may have any suitable cross-sectional shape,i.e., symmetrical or asymmetrical, but most often have a symmetricalcross-sectional shape, preferably a spherical, cylindrical or polylobalshape. The cross-sectional diameter of the particles may be as small as40 μm; however, it is usually about 0.635 mm (0.25 inch) to about 12.7mm (0.5 inch), preferably about 0.79 mm ( 1/32 inch) to about 6.35 mm(0.25 inch), and most preferably about 0.06 mm ( 1/24 inch) to about4.23 mm (⅙ inch). Among the preferred catalyst configurations arecross-sectional shapes resembling that of a three-leaf clover, as shown,for example, in FIGS. 8 and 8A of U.S. Pat. No. 4,028,227. Preferredclover-shaped particulates are such that each “leaf” of thecross-section is defined by about a 270° arc of a circle having adiameter between about 0.51 mm (0.02 inch) and 1.27 mm (0.05 inch).Other preferred particulates are those having quadralobalcross-sectional shapes, including asymmetrical shapes, and symmetricalshapes such as in FIG. 10 of U.S. Pat. No. 4,028,227.

Since the reaction is conducted under at least partial liquid phaseconditions, reaction pressure may be adjusted to maintain the ethyleneat least partially in the liquid phase. Ethylene in the gas phase mayalso be suitable. Pressures can vary within a wide range of about 101 toabout 13172 kPa (gauge) (1 to 1900 psig). As a practical matter thepressure normally is in the range between about 1379 and about 6985 kPa(gauge) (200 to 1000 psig) but usually is in a range between about 2069and 4137 kPa(gauge) (300 and 600 psig). The temperature rangeappropriate for alkylation of the benzene with the ethylene is betweenabout 100° and about 300° C. The ratio of aromatics to ethylene shouldbe between about 1:10 and as high as about 10:1, with a ratio of 0.5 to1.0 being preferred.

The dilute ethylene feed may be fed to the alkylation reactor 320 in thegas phase at a temperature between about 100° and about 300° C. Thereaction takes place predominantly in the liquid phase at a WHSV 0.01 to10 hr⁻¹ on ethylene basis. The gaseous ethylene may absorb into theliquid benzene stream for alkylation to occur. We have found,surprisingly, that despite the presence of impurities in the feed thatpoison the catalyst and dilute the ethylene in the feed, that at leastabout 40 wt-% and as much as about 75 wt-% of the ethylene in the feedstream alkylate with aromatic rings to convert to heavier alkylaromatichydrocarbons.

The dilute benzene feed may be fed to the alkylation reactor 320 in theliquid phase at a temperature between about 100° and about 300° C. Thereaction takes place predominantly in the liquid phase at a WHSV of 0.1to 40 hr⁻¹ on benzene basis. We have found, surprisingly, that despitethe presence of heavier aromatics and paraffins in the feed that dilutethe benzene in the feed, that at least about 20 wt-%, suitably at leastabout 50 wt-% and as much as about 100 wt-% of the benzene in the feedstream convert to heavier alkylbenzene. Moreover, the conversion ofbenzene is at least about 80%, preferably at least about 90% and mostpreferably at least about 95% of the conversion of toluene. Even thoughthe dilute benzene feed may have a higher concentration of aromaticshaving eight carbon atoms than the benzene concentration, the benzeneundergoes greater conversion than the aromatics having eight carbonatoms. The benzene to olefin ratio may be between 0.2 and 4.0.

The catalyst remains stable despite the impure feed, but it can beregenerated upon deactivation. Suitable regeneration conditions includesubjecting the catalyst, for example, in situ, to hot air at 500° C. for3 hours. Activity and selectivity of the regenerated catalyst iscomparable to fresh catalyst.

The alkylation product stream from the alkylation reactor in line 352can be transported to an alkylation product fractionation column 360which may be a simple flash drum but is preferably a fractionationcolumn to separate a gaseous stream from a liquid stream. The alkylationproduct fractionation column 360 is in downstream communication with thealkylation reactor 320. The gaseous product stream in overhead line 362comprising light gases such as hydrogen, methane, ethane, unreactedolefins and light impurities may be transported to a combustion unit 370to generate steam in line 372. Alternatively, the gaseous product inoverhead line 362 may be combusted to fire a heater (not shown) and/orto provide a source of flue gas to turn a gas turbine (not shown) togenerate power. The overhead line 362 is in upstream communication withthe combustion unit 370. The liquid bottoms stream comprising heavierhydrocarbons in line 364 from the alkylation product fractionationcolumn 360 can be let down over a valve and recycled back to the productseparation section 90 via LCO pump-around 95. Consequently, the maincolumn 92 is in downstream and upstream communication with thealkylation reactor 320. The bottoms stream in line 364 is preferablyrecycled to the main column 92 at a location between the heavy naphthaoutlet 96 a and the light cycle oil outlet 95 a. The recycle line 364 isin downstream communication with a bottom of the alkylation productfractionation column 360. Alternatively, the recycle line 364 feeds thelight cycle oil pump-around line 95 or the heavy naphtha pump-aroundline 96. The recycle line 364 is in downstream communication with thealkylation reactor 320 and in upstream communication with the maincolumn 92. Alternatively, the alkylation product in lines 364 may betransported to the gasoline distribution system 148 without recycling tothe product separation zone 90.

A stream of naphtha range alkyl benzene with a smaller concentration ofbenzene than in line 304 which may be from a side cut in outlet line 366may be recovered and delivered to the gasoline distribution system 148.The gasoline distribution system 148 may comprise piping to an outlet,to a dispenser for filling a tank for transportation or to a gasolinestorage tank. The gasoline distribution system 148 is in downstreamcommunication with the alkylation reactor 320. In an embodiment, theside cut stream in line 366 will have a larger flow rate than thebottoms stream in line 364, which may just be a drag stream.

EXAMPLES

The utility of the present invention will be demonstrated by thefollowing examples.

Example 1

An extruded UZM-8 catalyst was synthesized by combining an UZM-8 powderwith Si/Al ratio of 12, and pseudo-boehmite provided under the Catapaltrademark. The pseudo-boehmite was peptized with nitric acid beforemixture with the amorphous silica-alumina. The catalyst dough wasextruded through 1.59 mm openings in a cylindrical die plate and brokeninto pieces prior to calcination at 550° C. The finished catalystconsisted of 70 wt-% UZM-8 and 30 wt-% alumina and had a surface area of368 m²/g.

Example 2

The catalyst of Example 1 was tested for benzene alkylation at 205° C.,3447 kPa (500 psig), 0.4 OWHSV (olefin weight hourly space velocity) anda liquid feed WHSV of 4.9 in a fixed bed operation over 10 mL ofcatalyst. The gas feed consisted of 23 mol-% C₂H₄ and 77 mol-% CH₄. Theliquid feed consisted of 2.6 wt-% pentene, 6.4 wt-% C₆H₆, 30 wt-%n-heptane, 25 wt-% toluene, 18 wt-% aromatics having eight carbon atoms(C₈ aromatics), and 19 wt-% aromatics having nine carbon atoms (C₉aromatics).

Of ethylene, 83% was converted, and 51% of benzene was converted.Toluene conversion was 49% while conversion of C₈ and C₉ aromatics wasless than 20%. Selectivity to ethylbenzene was 5%, to ethyltoluenes was25%, to diethylbenzenes was 6%, to other gasoline range compounds was37% and to compounds boiling at greater than 225° C. was 27%.

After 13 hours, the temperature was increased to 236° C. At the 26 hourmark, the liquid feed WHSV was increased to 12.7. Under theseconditions, the ethylene conversion was 83%, with benzene conversion of32%, toluene conversion of 27% and C₈ and C₉ aromatic conversion of lessthan 15%. Selectivity was 4% to ethylbenzene, 17% to ethyltoluenes, 2.5%to diethylbenzenes, 53.5% to other gasoline range compounds and 23% tocompounds boiling at greater than 225° C.

At 46 hours on stream, the gas feed was changed to a blend representingan FCC dry gas composition containing 34 mol-% CH₄, 23 mol-% C₂H₄, 14mol-% C₂H₆, 13 mol-% H₂, 13 mol-% N₂, 2 mol-% CO₂, 1 mol-% CO and 1 ppmH₂S. Conversions and selectivities did not change. This experiment isshown in FIG. 2.

Example 3

The catalyst of Example 1 was tested for benzene alkylation at 205° C.,3447 kPa (500 psig), 0.8 OWHSV and a liquid feed WHSV of 7 in a fixedbed operation over 12 mL of catalyst. The gas feed consisted of 34 mol-%CH₄, 23 mol-% C₂H₄, 14 mol-% C₂H₆, 13 mol-% H₂, 13 mol-% N₂, 2 mol-%CO₂, 1 mol-% CO and 1 ppm H₂S resembling a dry gas feed. The liquid feedconsisted of 39 wt-% benzene, 49 wt-% n-heptane and 12 wt-% toluene.Ethylene conversion ranged from 77% to 58% over 50 hours on stream witha constant 44% benzene conversion. Toluene conversion was 47%.Selectivity was 50-53% to ethylbenzene, 14-16% to ethyltoluenes, 17-19%to diethylbenzenes, 11-14% to other gasoline range compounds and 3% tocompounds boiling at greater than 225° C. This experiment is shown inFIG. 3.

Without further elaboration, it is believed that one skilled in the artcan, using the preceding description, utilize the present invention toits fullest extent. The preceding preferred specific embodiments are,therefore, to be construed as merely illustrative, and not limitative ofthe remainder of the disclosure in any way whatsoever.

In the foregoing, all temperatures are set forth in degrees Celsius and,all parts and percentages are by weight, unless otherwise indicated.

From the foregoing description, one skilled in the art can easilyascertain the essential characteristics of this invention and, withoutdeparting from the spirit and scope thereof, can make various changesand modifications of the invention to adapt it to various usages andconditions.

1. An apparatus for alkylating benzene with ethylene comprising: a fluidcatalytic cracking reactor for contacting cracking catalyst with ahydrocarbon feed stream to crack the hydrocarbon feed to crackedproducts having lower molecular weight and deposit coke on the crackingcatalyst to provide coked cracking catalyst; a regenerator forcombusting coke from said coked cracking catalyst by contact withoxygen; a separator in communication with said fluid catalytic crackingreactor for separating C₃ hydrocarbons from C₂ hydrocarbons to provide adilute ethylene stream; a reforming reactor for contacting a naphthastream with reforming catalyst to produce a reformate stream; and analkylation reactor in communication with said separator and saidreforming reactor for alkylating benzene in said reformate stream withethylene in said dilute ethylene stream over a fixed bed of alkylationcatalyst to heavier alkyl benzene hydrocarbons, wherein no fractionationcolumn is in communication between said separator and said alkylationreactor.
 2. The apparatus of claim 1 wherein said separator is asecondary absorber and further comprising a primary absorber incommunication with said secondary absorber for providing a secondaryoff-gas stream comprising said dilute ethylene stream and saidalkylation reactor is in communication with said secondary absorber. 3.The apparatus of claim 1 wherein no fractionation column is incommunication between said reforming reactor and said alkylationreactor.
 4. The apparatus of claim 1 wherein a reformate splitter columnis in communication between said reforming reactor and said alkylationreactor.
 5. The apparatus of claim 4 wherein a line exiting a bottomhalf of the reformate splitter column bypasses the alkylation reactor.6. The apparatus of claim 4 wherein a line exiting a top half of thereformate splitter column communicates with the alkylation reactor. 7.The apparatus of claim 1 wherein a main fractionation column is incommunication with said FCC reactor and said separator is incommunication with said main fractionation column.
 8. The apparatus ofclaim 1 further comprising a gasoline distribution system incommunication with said alkylation reactor.
 9. An apparatus foralkylating benzene with ethylene comprising: a fluid catalytic crackingreactor for contacting cracking catalyst with a hydrocarbon feed streamto crack the hydrocarbon feed to cracked products having lower molecularweight and deposit coke on the cracking catalyst to provide cokedcracking catalyst; a cracked product outlet for discharging said crackedproducts from said reactor; a regenerator for combusting coke from saidcoked cracking catalyst by contact with oxygen; an absorber incommunication with said product outlet for providing an off-gas streamcomprising a dilute ethylene stream; a reforming reactor for contactinga naphtha stream with reforming catalyst to produce a reformate stream;and an alkylation reactor in communication with said absorber and saidreforming reactor for alkylating benzene in said reformate stream withethylene in said dilute ethylene stream over a fixed bed of alkylationcatalyst to heavier alkyl benzene hydrocarbons, wherein no fractionationcolumn is in communication between said reforming reactor and saidalkylation reactor and between said absorber and said alkylationreactor.
 10. The apparatus of claim 9 wherein said absorber is asecondary absorber and further comprising a primary absorber incommunication with said secondary absorber for providing a secondaryoff-gas stream comprising said dilute ethylene stream and saidalkylation reactor is in communication with said secondary absorber. 11.The apparatus of claim 9 further comprising a gasoline distributionsystem in communication with said alkylation reactor.
 12. The apparatusof claim 11 further comprising an alkylation product fractionationcolumn in communication with said alkylation reactor with an outlet linein communication with said gasoline distribution system.